Direct coal liquefaction process

ABSTRACT

A direct coal liquefaction method and apparatus in which the feed coal is mixed with a recycled 600° F.+ non-donor stream in which the ratio of coal to said stream is at least 1.5:1 on a moisture free basis to form an input slurry to a DCL reactor. Hydrogen containing treat gas is supplied to the reactor. 1000° F.− bottoms from the reactor are recycled as part of the 600° F.+ non-donor stream. 1000° F.+ bottoms from the reactor are gasified in a PDX unit to provide hydrogen for the DCL reaction. The ratio of recycled bottoms to feed coal is between 1:0.5 and 1:1.5.

FIELD OF THE INVENTION

The present invention relates to a direct coal liquefaction processcapable of producing gasoline, jet fuel, or diesel from high inertinitecontent coal at high thermal efficiency.

BACKGROUND OF THE INVENTION

Increases in the cost of petroleum and concerns about future shortageshave led to increased interest in the use as a fuel source of the vast,easily accessible deposits of coal that exist in several parts of theworld. Various processes have been proposed for converting coal toliquid and gaseous fuel products including gasoline, diesel fuel,aviation turbine fuel and heating oil, and, in some cases, to otherproducts such as lubricants and chemicals. A number of problems havehampered widespread liquefaction of coal, however, including therelatively low thermal efficiency of indirect coal-to-liquids (CTL)conversion methods, such as Fischer Tropsch (FT) synthesis andmethanol-to-liquids (MTL) conversion, and concerns about CO₂ emissions.Direct coal liquefaction (DCL) methods have been developed forliquefying coal that have advantages in many applications relative toconversion by FT synthesis, including substantially higher thermalefficiency and lower CO₂ emissions. Such direct liquefaction methodshave typically involved heating the coal in the presence of a donorsolvent, and optionally a catalyst, in a hydrogen containing atmosphereto a temperature in the range of about 700° to 850° F. to break down thecoal structure into free radicals that are quenched to produce liquidproducts. The catalyst can typically be very finely divided iron ormolybdenum or mixtures thereof. One source of the molybdenum catalyst isvia in situ formation from a phosphomolybdic acid (PMA) precursor.

The reported DCL coal conversion units require the use of a hydrotreaterfor preparing the donor solvent that is fed back to the input of the DCLunit to act as a solvent for the coal being converted and to provideadditional hydrogen to the liquefaction process. Solvent hydrotreatingrequires separation of the solvent fraction, additional equipment,additional hydrogen-rich treat gas, and as much as 50% of the heat ofreaction is moved from the liquefaction reactor to the solventhydrotreater. Thus, addition of an external solvent hydrotreaterincreases the required investment and decreases thermal efficiency. Thisis an important reason why low donor solvent-to-coal ratios (typically1.2 to 1.5) are used in processes using a solvent hydrotreater. Inaddition, the hydrotreating reduces the viscosity and lowers thearomatics content of the solvent, which reduces its ability to suspendash in the slurry and its compatibility with coal. The reduction inability to suspend ash results in an increased likelihood of solidsbuildup, deposits, or plugging of high pressure feed pumps, transferlines, heat exchangers, furnace tubes, and reactors. Therefore, use ofsuch donor solvents results in higher gas hold-up in the liquefactionreactor in order to maintain the solids in the slurry in suspension,which in turn, requires a large reactor volume to achieve adequate coalresidence time in the reactor. High recycle gas rates are also requiredbecause treat gas must be provided to both the solvent hydrotreater andthe liquefaction reactor. Because hydrotreaters are expensive, requireenergy for recycling gas, and use the heat of reaction less efficiently,DCL reactor systems have been designed to minimize the amount of donorsolvent recycle.

In order to reduce the required reactor volume and minimize solidsbuild-up in the liquefaction reactors, Shenhua and Headwaters utilizeebullated bed reactors. Circulation of liquid through the reactorsresults in a decrease in gas hold-up. Because of the high liquid recyclein ebullated bed reactors, the reactors are fully back-mixed whichresults in an increase in reactor volume versus a plug flow reactor.

A further issue limiting the application of DCL methods is that lowerquality coals having inertinite content much higher than about 12% havebeen considered unsuitable for use as a DCL feed stock. Such highinertinite coals are found in many parts of the world, including theUnited States and China. Many of these coals, such as that in the Ordosbasin in China (1,2,3), have inertinite content of more than 25% and alow ratio of atomic hydrogen to carbon (H/C) and have historically beenunacceptably more difficult to liquefy by DCL than higher quality coalsthat have a high vitrinite content.

In 1995, Okada (4) reported that oil yield from autoclave experimentswere inversely proportional to the inertinite content of coal (FIG. 2)for coals of similar rank. Oil yield ranged from 67 wt % for a zeroinertinite content coal to 40 wt % for a coal containing 60 vol %inertinite. He concluded that high inertinite coals, such as found inthe Ordos Basin, are not suitable for direct liquefaction.

In 2001, Wasaka (5) (NEDOL) published results on 53 runs on 27 coalsthat were made in a 0.1 t/d pilot plant test program. The programspecifically focused on identifying the preferred conditions forliquefying Chinese coals, including high inertinite coals. The studyconcluded that a combination of iron catalyst and an externallygenerated donor solvent produced the highest oil yield, particularly fora high inertinite content coal (FIG. 3). At 450° C., 17 MPa, donorsolvent, and FeS2 catalyst, an oil yield of approximately 46 wt % wasobtained at a pilot plant recycle gas treat rate of 13 Nm3/hr.Increasing recycle gas rate to the reactor (excludes treat gas requiredfor solvent hydrotreating) resulted in an increase in oil yield toapproximately 60 wt % (FIG. 4).

Based on these experiments, Wasaka concluded that obtaining higher oilyields was difficult in liquefaction of high inertinite content coalsbut that an iron catalyst, a donor solvent, temperatures up to 460° C.,and a high recycle gas rate were preferred for maximizing oil yield.Ishibashi (6,7), et. al., published hydrodynamic information on a seriesof slurry liquefaction reactors at varying operating conditions. All ofthe operations utilized an externally hydrotreated donor solvent. Gashold-up in the slurry reactors increased linearly with superficialvelocity to about 8 cm/sec and then leveled out at approximately 70percent. This line out is typically associated with a reactor that movesfrom bubbly flow to churned turbulent operation. This high gas hold-upand corresponding lower solid and liquid hold-up results in an increasein liquefaction reactor volume required to achieve a given coalresidence time and coal conversion to liquid products.

In 2005, Shenhua (7,8) applied for a patent (issued in 2010) for a DCLprocess that utilized iron catalyst and an externally hydrotreated donorsolvent. The donor solvent is produced in a suspended bed with a forcedcirculation reactor (ebullated bed). Suspended bed reactors with forcedcirculation were selected for liquefaction to provide lower gas hold-up(high utilization of reactor volume) and to avoid precipitation ofmineral salts. Two suspended beds reactors with forced circulation areutilized for liquefaction with a treat gas rate (0.6-1.0 Nm3/kg slurry).The preferred embodiment operated both liquefaction reactors at 455° C.,a reactor pressure of 19 MPa, a solvent to coal (S/C) ratio of 1.2, andan iron catalyst addition rate of 1 wt % Fe on dry coal. Oil yield for alow ranked bituminous coal is 57.17 wt % on a moisture and ash free(MAF) basis. In a separate publication, Shenhua reported ThermalEfficiency for DCL as 59.75%.

The Headwaters DCL Process(9,10) features include use of disperse, ironbased catalyst, two-stage, back mixed, slurry phase reactors, mildhydrogenation to replenish the hydrogen content of the recycle solventand to stabilize the raw distillate products, and deashed residconversion to enhance net coal conversion. Reactor temperature in theliquefaction reactors is 435° C. to 460° C. at 17 Mpa.

In both the Shenhua and Headwaters processes, products would be furtherupgraded using conventional refining processes.

SUMMARY OF THE INVENTION

In accordance with the invention, it has been found that a dramaticincrease in liquid product yield and thermal efficiency is achieved whenthe solvent to coal ratio is increased in a DCL process, particularly inmicrocatalytic coal liquefaction (MCL) processes that employ very finelydivided molybdenum or iron catalysts, preferably a finely dividedmolybdenum catalyst, and a 600 to 700° F.+ non-donor recycle streamproduced in liquefaction. The ratio of such stream to coal at the inputto the reactor on a moisture free weight basis is greater than 1.6:1,preferably greater than 1.7:1, more preferably between 1.8:1 and 3.5:1,still more preferably between 2.0:1 and 3.5:1, and most preferablybetween 2.0:1 and 3.0:1. By “non-donor” is meant that the recycle streamhas not been processed in a hydrotreater to partially hydrogenatemulti-ring aromatic compounds in the stream to produce compounds thatcan donate hydrogen during liquefaction.

Surprisingly, increasing the ratio of the recycled 600 to 700° F.+non-donor stream to coal in the MCL process does not increase the flowrate of recycled stream and fresh coal to liquefaction for a given rateof product generation. Instead, less coal is required and although therecycle stream increases relative to coal, the total feed toliquefaction remains essentially the same. The net impact of higherrecycle and lower coal rate is a reduction of energy required in theslurry preheat furnace and the size of the vacuum fractionator. Hence,investment and energy requirements are reduced for the liquefactionsection of the MCL plant.

Moreover, the MCL process of the invention has been found to efficientlyliquefy even coals having inertinite contents much higher (e.g., morethan 25%) than were previously thought to be suitable as feeds fordirect coal liquefaction processes.

Because the viscosity and density of the coal slurry in the process ofthe present invention is much higher than is the case with slurry usinga hydrotreated donor solvent, the maintenance of a stable slurry and thesettling of ash in the reactor is not a concern. Therefore the processof the invention operates with a much lower gas hold-up than is requiredin hydrotreated donor solvent DCL systems. Finally, volume occupied byinternals in the ebullated bed reactors is eliminated. As a result ofthe low gas hold-up and the preferred use of slurry reactors in series,the reactor volume can be significantly less than that required for aDCL system having the same output capacity operating with a donorsolvent and high gas hold-up. Moreover, the use of a non-donor 600 to700° F.+ stream eliminates the need for a hydrotreater for hydrotreatingthe donor solvent, which substantially reduces the complexity andcapital cost of the system.

The MCL process of the invention has been found to efficiently liquefyeven coals having inertinite contents much higher than were previouslythought to be suitable as feeds for direct coal liquefaction processes.Moreover, it has been found that an MCL plant operating in accordancewith the process of the invention can produce easily upgraded liquidsand operate at a high thermal efficiency, in the range of approximately65% to over 70%, even when processing high inertinite coal feeds.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a graph of the ratio of hydrogen to carbon vs. inertinitecontent for a variety of U.S. and Chinese coals.

FIG. 2 is a schematic diagram of a direct coal liquefaction systemsuitable for use in the illustrated embodiment of the invention.

DETAILED DESCRIPTION OF ILLUSTRATED EMBODIMENT

The DCL process of the invention achieves a higher conversion of coal toliquid fuels and a higher thermal efficiency than those of other directcoal liquefaction processes, even with coals having high inertinitecontents that were previously considered unsuitable as a feed for directcoal liquefaction processes. As illustrated in FIG. 1 of the drawings,coals having higher inertinite content tend to have lower ratios ofhydrogen to carbon (H/C). As shown, in the extremes, the H/C ofvitrinite and exinite coals is approximately 0.85. Extrapolation of theregression line to approximately 100% inertinite indicates the H/C ofthe inertinite is 0.5. This low H/C reflects a coal structure that has ahighly condensed aromatic ring structure that is very difficult toliquefy.

Referring now to the embodiment of a DCL system illustrated in FIG. 2 ofthe drawings, the coal feed is dried and crushed in a conventional gasswept roller mill 201 to a moisture content of 1 to 4%. Crushed anddried coal is fed into a mixing tank 203 where it is mixed with a streamconstituted by a 600 to 700° F.+ fraction, preferably a 650° F.+fraction, of the output of the liquefaction reactor to form a slurrystream. The catalyst precursor in the illustrated embodiment preferablyis in the form of an aqueous water solution of phosphomolybdic acid(PMA) in an amount that is equivalent to adding between 50 wppm and 2%molybdenum relative to the dry coal feed. In the slurry mix tank 203,typical operating temperature ranges from 300 to 600° F. and morepreferably between 300 and 500° F. From the slurry mix tank, thecatalyst containing slurry is delivered to the slurry pump 205. Theselection of the appropriate mixing and temperature conditions is basedon experimental work quantifying the rheological properties of thespecific slurry blend being processed.

Most of the remaining moisture in the coal is driven off in the mixingtank due to the hot atmospheric fractionator bottoms feeding to themixing tanks. Residual moisture and any entrained volatiles arecondensed out as sour water (not shown in FIG. 2). The coal in theslurry leaving the mixing tank 203 has about 0.1 to 1.0% moisture. Theslurry formed by the coal, 600 to 700 to 1,000° F. stream from thevacuum fractionator 221, and the 600 to 700° F.+ stream fraction fromthe atmospheric fractionator 219 (also referred as an atmospheric pipestill or APS), is pumped from the mixing tank 203 and the pressure israised to about 2,000 to 3,000 psig (138 to 206 kg/cm² g) by the slurrypumping system 205. The resulting high pressure slurry may be preheatedin a heat exchanger (not shown), mixed with a treat gas consisting ofrecycled and makeup treat gas containing over 80% hydrogen, and thenfurther heated in furnace 207.

The coal slurry and hydrogen mixture is fed to the input of the firststage of the series-connected liquefaction reactors 209, 211 and 213 atbetween 600 to 700° F. (316 to 371° C.) and 2,000 to 3,000 psig (138 to206 kg/cm² g). The reactors 209, 211 and 213 are simple up-flow tubularvessels, the total length of the three reactors being 40 to 200 feet.The temperature rises from one reactor stage to the next as a result ofthe highly exothermic coal liquefaction reactions. In order to maintainthe maximum temperature in each stage below about 800 to 900° F. (427 to482° C.), a portion of the hydrogen based treat gas is preferablyinjected between reactor stages. The hydrogen partial pressure in eachstage is preferably maintained at a minimum of about 1,000 to 2,000 psig(69 to 138 kg/cm²g).

The effluent from the last stage of liquefaction reactor is separatedinto a gas stream and a liquid/solid stream, and the liquid/solid streamlet down in pressure, in the separation and cooling system 215. The gasstream is cooled to condense out the liquid vapors of H2O, naphtha,distillate, and solvent. The remaining gas is then processed to removeH₂S and CO₂

Most of the processed gas is then sent to a hydrogen recovery system,not shown, for further processing by conventional means to recover thehydrogen contained therein, which is then recycled to be mixed with thecoal slurry. The remaining portion of the processed gas is purged toprevent buildup of light ends in the recycle loop. Hydrogen recoveredtherefrom can be used in the downstream hydro-processing upgradingsystem.

The depressurized liquid/solid stream and the hydrocarbons condensedduring the gas cooling are sent to the atmospheric fractionator 219where they are separated into light ends, naphtha, distillate, and 600to 700° F.+ fractions. The light ends are processed to recover hydrogenand C₁-C₄ hydrocarbons that can be used for fuel gas and other purposes.The naphtha is hydrotreated to saturate olefins and other reactivehydrocarbon compounds. The 160° F.+ fraction of the naphtha can behydrotreated and catalytically reformed to produce gasoline. Thedistillate fraction can be hydrotreated to produce products such asdiesel and jet fuel.

A portion of the 600 to 700° F.+(316 to 371° C.+) is recycled to theslurry mix tank. The remaining 600 to 700° F.+ fraction produced fromthe atmospheric fractionator 219 is fed to the vacuum fractionator 221(also referred to as a vacuum pipe still) wherein it is separated into a1000° F.− fraction and a 1000° F.+ fraction. The 1000° F.− fraction isadded to the 600 to 700° F.+ stream being recycled to the slurry mixtank 203.

In the illustrated embodiment, the 1000° F.+ fraction from the vacuumfractionator 221 is sent to be gasified by the partial oxidation system223 to generate hydrogen for use in the liquefaction. A portion of thecoal from the gas swept mill 201 is also fed to the partial oxidationsystem 223 to produce additional hydrogen. Alternatively, instead of thepartial oxidation system 223, the 1000° F.+ bottoms from the vacuumfractionator 221 may be processed in a Circulating Fluid Bed boiler, acement plant, or sold as a feed for asphalt paving or electrodemanufacture. G.E., Shell, and others offer commercial processes forgasification (partial oxidation) of the 1000° F.+ bottoms andCirculating Fluid Bed boiler manufactures such as Foster-Wheeler andAlstom offer technology for combusting the 1000° F.+ bottoms.

Hydrogen for liquefaction and upgrading can also be produced by SteamMethane Reforming of a stream such as natural gas, shale gas, or coalmine methane. This technology is utilized worldwide in refineries andoffered by many commercial vendors such as Haldor-Topsoe.

Catalysts useful in DCL processes also include those disclosed in U.S.Pat. Nos. 4,077,867, 4,196,072 and 4,561,964, the disclosures of whichare hereby incorporated by reference in their entirety.

Other DCL reactor systems suitable for use in the process of theinvention are disclosed in U.S. Pat. Nos. 4,485,008, 4,637,870,5,200,063, 5,338,441, and 5,389,230, and U.S. patent application Ser.No. 13/657,087, the disclosures of which are hereby incorporated byreference in their entirety.

The preferred DCL Process combines several elements that contribute tomaximum Premium Fuels Product production and maximum thermal efficiency.These include, very importantly, the recycle of a non-donor 600 to 700°F.+ stream, preferably including atmospheric fractionator bottoms, tomaintain a ratio of the recycle stream to coal at the input to thereactors 209, 211, 213 that is greater than 1.6:1 on a moisture freeweight basis, preferably greater than 1.7:1, more preferably between1.8:1 and 3.5:1, still more preferably between 2.0:1 and 3.5:1, and mostpreferably between 2.0:1 and 3.0:1; the use of a microcatalyst in theform of finely divided molybdenum; and the use of a much lower treat gasrate than in previous systems. Also, use of bottoms recycle, andmultiple slurry reactors in series contribute to the benefits of theprocess.

-   -   (1) Use of a microcatalyst, which is either a compound of        molybdenum or iron, more preferably molybdenum, and added at 100        to 1,000 wppm, more preferably 100 to 500 wppm, and most        preferably 100 to 300 wppm, eliminates several disadvantages to        the use of a donor solvent such as required by prior DCL        systems. First, energy is lost during preparation of the donor        solvent. Second, energy is required to preheat the donor solvent        in the solvent hydrotreater and hydrogen must be compressed and        circulated around the hydrotreater. Thirdly, the heat release        during partial hydrogenation of the donor solvent is lost during        cooling prior to separation of hydrogen for recycle. In        comparison, all of the heat release occurs in the liquefaction        reactors during operation with a 600 to 700° F.+ recycle stream,        which minimizes the preheat requirement prior to liquefaction.        These factors contribute to the higher thermal efficiency of the        microcatalytic coal liquefaction process. Moreover, the use of a        microcatalyst and the consequent elimination of the need for a        donor solvent also eliminates the need for an expensive solvent        hydrotreater to generate the donor solvent, thereby        substantially reducing the capital cost of the system. It also        permits the use of coals having substantially higher ash        contents, from 6 to 20 wt % or more on a moisture free basis,        and the recycle of a substantially higher portion of bottoms        than were possible with donor solvent systems. Examples of        microcatalysts and their method of preparation are described in        U.S. Pat. No. 4,226,742, the contents of which are hereby        incorporated by reference in their entirety.    -   (2) The 600 to 700° F.+ fraction recycled from the atmospheric        fractionator 219 and the 1000° F.− fraction from the vacuum        fractionator 221 as the non-donor stream being recycled to the        slurry mix tank 203 provides preheat for the coal and solvent in        the slurry mix tank 203. This raises the temperature in the mix        tank to 200° F. to 500° F., more preferably 350° F. to 500° F.,        and most preferably about 400 to 500° F. This further reduces        the energy requirement for preheating the slurry prior to        liquefaction. A significant portion of the of the microcatalyst        is entrained in the 600 to 700° F.+ fraction recycled from the        atmospheric tower 219, so that recycling a larger portion of        such fraction increases the catalyst concentration in the DCL        reactors 209, 211, 213, thereby decreasing the requirement for        the addition of fresh catalyst precursor and increasing the        conversion efficiency of the DCL process.    -   (3) Use of the non-donor 600° F. to 700° F.+ stream, more        preferably 630° F. to 670° F.+, and most preferably a 650° F.+,        process derived recycle solvent in the DCL process reduces        cracking, relative to donor solvent, and produces a 650° F.−        product with a greater fraction of diesel and less light gases        and naphtha. The 650° F.− product can be selectively upgraded to        finished products in fixed bed upgrading reactors.    -   (4) The much lower treat gas rate of 600 to 900 NL per kg of        slurry has a significant impact on thermal efficiency, plant        investment, and operating cost. The required recycle treat gas        rate for the DCL process of the invention is up to three times        lower than the preferred gas rate in the NEDOL program (without        taking into account the treat gas rate to the solvent        hydrotreater, which makes the difference even larger). This has        an important impact on power requirements for the compressor and        fuel requirements for slurry preheat furnace 207 and solvent        hydrotreater preheat.    -   (5) The use of two to four, more preferably three slurry        reactors in series approaches a plug flow reactor and hence has        as little as two thirds of the required volume of one or two        ebullated bed reactors such as used in some prior DCL systems.        Since all of the heat is released in the three liquefaction        reactors, the temperature profile can be also maintained to        maximize selectivity to liquids. Operation of the initial        reactor at a somewhat lower temperature has been reported in        previous patents as a route to increase conversion and liquid        yields.

An exemplary process for upgrading the liquid product of the DCLreactors 209, 211, 213 is disclosed in U.S. Pat. No. 5,198,099, thedisclosure of which is hereby incorporated by reference in its entirety.Other processes and systems suitable for upgrading the liquid productsare commercially available from vendors such as UOP, Axens, Criterionand others.

The diesel product after upgrading will have a Cetane number of betweenapproximately 42 and 47, depending upon cut points of the product andaromatics content. Specific gravity of the product will also varybetween 0.83 and 0.90. A higher Cetane Number is required for Euro 4diesel, thus a Fischer-Tropsch facility producing a 70-75 Cetane Numberdiesel blend stock may be added to the plant operating in accordancewith the present invention. The gasoline produced by upgrading therelevant portion of product of the process of the present invention willmeet all current gasoline specifications, or can be upgraded to aResearch Octane of 106 if desired. This will permit the blending of thelow octane naphtha into the gasoline pool while maintaining adequateoctane for the blended fuel.

The upgrading process can also be operated to maximize the production ofjet fuel or gasoline. The jet fuel produced will meet all Military JP-8specifications.

Example

The process of the invention was run in the same 0.1 t/d pilot plantused in the runs described by Wasaka. Modifications were made to thepilot plant to the slurry mix tank, pumps, slurry reactors,microcatalyst (moly) addition, and 600 to 700° F.+ recycle. Compared toWasaka's pilot plant operations, slurry feed rate was increased to 10 to12 kg/hour and gas rate was reduced by a factor of 2 to 4. During thepilot plant operation, the reactors did not experience solids buildup orreactor deposits. The coal used in this example is shown as the squaresymbol in the chart of FIG. 1 and is therefore consistent with overalltrend of H/C versus inertinite content.

Oil yield at a 1.5/1 ratio of non-donor 650° F.+(343° C.) recycle streamto coal was 44.6 wt % (MAF basis) When the 650° F.+ stream rate wasincreased to 3/1 at a low H2 feed rate of 0.73 Nm3/kg of slurry, theC5+/371° C. (C5/650° F.) yield increased to 66.4 wt %. This issubstantially higher than observed with an iron catalyst, donor solvent,and high recycle gas rate in a back mixed reactor. The oil yield is alsosignificantly higher than reported by others (U.S. Pat. No. 7,763,167B2.)

The net product from the high 650° F.+ recycle to coal run wasessentially free of 650° F.+. The lower boiling product was upgraded tofinished jet fuel and diesel in a conventional fixed bed upgrading unitand so avoiding the need for upgrading in an ebullated bed reactor inboth a solvent hydrotreater and a conventional upgrader. This not onlyeliminates the solvent hydrotreater but also the need to recycle treatgas to this unit.

The coal used in the two runs described above was a high inertinitecontent coal from the Ordos Basin with an H/C of 0.72.

Elemental Analysis Wt %, DAF Petrographic Analysis Carbon 81.58 Hydrogen4.92 Vitrinite, wt % 66.2 Nitrogen 1.04 Semi-vitrinite, wt % — Sulfur0.64 Micronite, wt % 0.4 Oxygen (by diff) 11.81 Inertinite, wt % 29.9Total 100 Ash, wt % dry coal 5.34

The following table presents operating conditions and oil yields forhigh inertinite coals from the Wasaka publication and FIGS. 3 and 4. Thefirst two columns present data from FIG. 4. The third column presentsthe highest data point for the 0.72 H/C data set which represents theShangwan coal which has an inertinite content of 31.5 vol %.

Source NEDOL NEDOL NEDOL NEDOL Figure FIG. 8 FIG. 8 FIG. 3 Coal InertContent, vol % 46.1 45.1 31.5 H/C 0.69 0.69 0.72 S/C/B 1.5/1/0 1.5/1/0n.a. Solvent Donor Donor Donor Catalyst FeS2 FeS2 FeS2 Make Up Catalyst,on coal n.a. n.a. n.a. Reactor CSTR CSTR CSTR Temperature, ° C. 450 450450 Pressure, Mpa 17 17 17 Slurry, kg/hr 8-10 8-10 8-10 Recycle Gas(Nm3/hr) 13 22 n.a. % H2 in Recycle Gas 90 90 H2 Makeup, Nm3/hr 5 5Total H2 to liquefaction, Nm3/hr 16.7 24.8 Liquids, wt % MAF 46 60 62

Even at a H2 gas rate of 24.8 Nm3/hour, the oil yield was onlyapproximately 60 wt % (MAF basis) for the data presented in FIG. 8 ofthe Wasaka article for a coal with an inertinite content of 46.1 vol %.Also included is the highest data point for the 0.72 H/C Shangwan coalplotted in FIG. 3. The yield for this case was approximately 62 wt %.S/C/B and recycle gas rate are not reported for this point.

It should also be noted that liquefaction yield in the Wasaka report isgiven only as oil yield. No information is provided on whether the oilyield is by distillation and/or extraction and the boiling range of theoil yield.

In comparison, operating conditions and oil yields are reported for the29.9% inertinite content coal described in the Example above.

Coal Inert Content, vol % 29.9 29.9 343° C.+/Coal 1.5/1 3.0/1 Solvent650° F.+ 650° F.+ Catalyst Mo Mo Make Up Catalyst, on coal 300 wppm 300wppm Reactor Slurry Slurry Reactor Temperature, ° F. 425/450/450420/450/450 Pressure, Mpa 17.5 17.5 Slurry, kg/hr 12 10 H2 toliquefaction (Nm3/hr) 6.0 7.3 C5/343° C. Liquids, wt % MAF 43.9 66.4

At a 343° C.+(650° F.+) fraction to coal ratio of 1.5, oil yield for theabove data is comparable to the first column of the NEDOL data. Itshould be noted that donor solvent and iron catalyst have been replacedwith 300 wppm of Microcatalyst. In addition, the gas rate toliquefaction has been reduced by almost a factor of 2.5. This excludesthe additional H2 that would be required for the solvent hydrotreater inthe NEDOL case.

The surprising result is presented in column 2 of the above results(operating in accordance with the process of the invention) versus thesecond and third columns of the NEDOL data. Increasing the 650° F.+solvent fraction to coal ratio from 1.5/1 to 3.0/1 resulted in anincrease in C5/343° C. (C5/650° F.) yield from 43.9 to 66.4 wt %. Thisis higher than either result reported in the Wasaka article and, as wasreported earlier, the 66.4 wt % oil yield from liquefaction is for aproduct boiling between C5 and 650° F., whereas the makeup of the yieldin Wasaka is not specified. More importantly, the H2 rate toliquefaction is only about 29% of the H2 rate required in the datapresented by Wasaka in their FIG. 8. This excludes the H2 rate requiredfor the solvent hydrotreater to produce donor solvent.

Hence, the recycle of solvent and hot bottoms from fractionationdirectly back to the slurry mix tank results in the highest recorded oilyield for a high inertinite coal while avoiding the need forcompression, heating, and cooling of over three times the H2 raterequired in the Wasaka operation.

The following cases were developed to demonstrate the impact of a highS/C ratio on the equipment in the liquefaction section of a commercialscale MCL DCL plant. The particular cases were developed for a 71 KB/Dplant with 650° F.+ fraction to coal ratio of 1.5/1 and 2.5/1. Coal andhydrogen requirement for liquefaction were estimated based on thereported pilot plant data.

It is interesting to note that the total quantity of feed from theslurry mix tank is constant for the two cases because of the decreasedamount of feed coal to liquefaction for the higher 650° F.+/Coal case.

Liquefaction Section Heat and Material Balance

650° F.+/Coal 1.5/1 2.5/1 C3-650 F. Product Rate, KB/D 71 71 % MAP Coalconversion to 61 83 650° F.− Coal to liquefaction, KT/D 21.5 15.4 Slurrymix tank, KT/D Coal 21.5 15.4 650/1000° F. from VPS 9.9 9.2 650° F.+From APS 22.4 29.2 Total 53.8 53.8

Since the product and selectivity are similar, the hydrogen consumptionand associated heat release are essentially the same.

Temperature rise in the liquefaction reactors is also equivalent in thetwo cases sense mass flow and heat release are approximately constant inthe liquefaction reactors. Because of the higher boiling point of the650° F.+ stream it remains liquid in the effluent from the liquefactionreactors.

After product separation, the liquid stream from the product separatorsflows to the Atmospheric Fractionator. 650° F.− product is distilledoverhead and part of the 650° F.+ product is recycled to the slurry mixtank.

Recycle from APS to slurry mix tank, KT/D

650° F.+ 22.4 29.2

Part of the 650° F.+ stream is sent to the Vacuum Fractionator where the650/1000° F. solvent fraction is recovered and recycled to the solventmix tank and the 1000° F.+ fraction leaves the liquefaction section. Asmentioned previously, this stream can be used in partial oxidation(PDX), in a Circulating Fluid Bed boiler, production of cement, etc.

The fraction that is recycled from the Vacuum Fractionator is shown inthe following table.

Recycle from VPS to slurry mix tank, KT/D

650/1000° F. Solvent 9.9 9.2

For this example, the temperature of the slurry mix increases for thehigh 650° F.+/Coal case and thus reduces the required duty for theslurry preheat furnace.

Thus the overall impact on equipment is shown in the following table.

Relative Slurry Preheat Furnace Duty 1.0 0.86 Relative Heat Release inliquefaction 1.0 1.0 Relative APS Diameter 1.0 1.0 Relative VPS Diameter1.0 0.93

Thus, the cost of equipment required for the high 650° F.+/Coal case isactually less than in the lower 650°+/Coal case. Thus investment shouldbe lower for the high 650° F.+/Coal option.

As indicated in the table below, the calculated thermal efficiency ofthe MCL liquefaction process for the two cases is substantially higherthan has been achievable in prior coal liquefaction processes.

Thermal Efficiency 79% 83%

The overall thermal efficiency for the a balanced MCL plant includinghydrogen generation, coal liquefaction, upgrading, and 1000° F.+ bottomsdisposal was calculated for plant's operating in accordance with theprocess of the invention with and without natural gas import. Coalconversion was reduced to 83% to maintain ash content to PDX of 25 wt %.

Case Coal Only Coal & Nat Gas Coal In, KST/SD 21.22 15.4 O2 In, KST/SD6.2 2.7 Coal Conversion, % 83 83.0 Ash Content to POX, wt % 24.9 24.9Heat In, GBtu/SD Coal 571.0 409.6 Natural Gas 0 107.5 Total In 571.0517.1 Heat Out Premium Fuels 379.9 379.9 Sulfur 0.9 0.7 Total Out 380.8380.6 TE, % HHV 66.7 73.6

As shown, thermal efficiency for a balanced, microcatalytic direct coalliquefaction plant (no import or export of power or gas) producingfinished gasoline, jet, or diesel from coal only is 66.7%. If naturalgas (or equivalent) is available, the thermal efficiency increases to73.6% This is a direct result of a higher solvent recycle, a majorreduction in recycle treat gas rate, the elimination of the solventhydrotreater, the decrease in the size of the slurry preheat furnace,and the release of the heat of reaction in the liquefaction reactorsrather than in the solvent hydrotreater. This thermal efficiency ishigher than any known thermal efficiency for a DCL plant operating on ahigh inertinite content coal such as present in the Ordos Basin.

REFERENCES

-   1. Weihua, A. “Coal Petrography and Genesis of Jurassic Coal in the    Ordos Basin, China”, Geoscience Frontiers, China University of    Geosciences (Beijing) Jul. 20, 2011-   2. Sitian Li, “Coal Resources and Coal Geography in China”,    Episodes, Vol. 18, mos. 1 & 2, 1995-   3. Peng Chen, “Petrographic Characteristics of Chinese Coals and    Their Application in Coal Utilization Processes”, Fuel 81 (2002)    1389-1395-   4. Okada, K., Possible Impacts of Coal Properties on the Coal    Conversion Technology, Coal Science, J. A. Pajares and J. M. D.    Tascon, 1995 Elsevier Science-   5. Wasaka, S., “Study on Coal Liquefaction Characteristics of    Chinese Coals”, Fuel 81 (2002) 1551-1557-   6. Ishibashi, H., “Gas Hold-up in Slurry Bubble Column Reactors of a    150 t/d Coal Liquefaction Pilot Plant Process”, Fuel 80 (2001)    655-664-   7. Zhang, Y., et. al., “Process for Direct Coal Liquefaction”, U.S.    Pat. No. 7,763,167 B2, Jul. 27, 2010-   8. Zhang, Y., Shenhua Group, Shenhua Coal Conversion Technology and    Industry Development-   9. Lee, T. L. K., “Status of Coal to Liquids Project Technology”,    McIIvaine Hot Topic Hour-Nov. 4, 2010-   10. R. Bauman, et. al., Direct Coal Liquefaction Process, U.S.    Provisional Application 61/553,981

1) A method for producing liquids from feed coal, comprising the stepsof: a) mixing a feed coal having an greater than 12% inertinite contentwith a non-donor stream that includes a recycled 600° F.+ fraction ofthe product from a direct coal liquefaction (DCL) slurry reactor to forma feed coal slurry, the ratio of non-donor stream to feed coal beinggreater than 1.6:1 on a moisture free basis; b) supplying said feed coalslurry to said DCL slurry reactor; c) adding a molybdenum or ironcontaining microcatalyst to the input of the DCL reactor; and d)supplying H2 containing treat gas to the DCL reactor. 2) The method ofclaim 1 wherein said microcatalyst consists essentially of molybdenum.3) The method of claim 2 wherein the concentration of the microcatalystadded during steady-state operation of the DCL reactor is equivalent to100 to 300 wppm relative to the feed coal on a moisture and ash free(MAF) basis. 4) The method of claim 1 wherein the ratio of saidnon-donor stream to feed coal is at least 1.7:1. 5) The method of claim1 wherein the ratio of said non-donor stream to feed coal is between1.8:1 and 3.5:1. 6) The method of claim 1 wherein the ratio of saidnon-donor stream to feed coal is between 2.0:1 and 3.0:1 7) The methodof claim 1 wherein said non-donor stream includes a portion of the 1000°F.+ fraction of the DCL product. 8) The method of claim 1 wherein theDCL reactor includes a plurality of series connected slurry liquefactionreactors. 9) (canceled) 10) The method of claim 1 wherein said feed coaland said non-donor stream are mixed in a slurry mix tank and thetemperature of the slurry in said slurry mix tank is between 300 and600° F. 11) The method of claim 10 wherein the temperature of saidslurry in said slurry mix tank is between 300 and 500° F. 12) (canceled)13) The method of claim 2 wherein said molybdenum microcatalyst is addedat a rate between 100 and 1,000 ppm by weight with respect to the coalfeed on an MAF basis. 14) The method of claim 2 to wherein saidmolybdenum microcatalyst is added at a rate between 100 and 500 ppm byweight with respect to the coal feed on an MAF basis during steady-stateoperation. 15) method of claim 1 where the ash content of the coal isbetween 6 and 20 wt % on the moisture free basis. 16) (canceled) 17) Themethod of claim 1 wherein the inertinite content of the coal is greaterthan 25%. 18) The method of claim 1 where hydrogen for liquefaction andupgrading is produced in a Steam Methane Reformer or an AutothermalReformer and 1000° F.+ fraction from DCl reactor is used for powergeneration, steam generation, or sold. 19) The method of claim 1 whereinthe ratio of treat gas to slurry supplied to the DCL reactor is lessthan 900 NL/kg of slurry.